Fluid hydroforming with catalyst recycle



April 10, 1956 A. L. CONN FLUID HYDROFORMING WITH CATALYST RECYCLE FiledDeo. 1 1951 FLuE @As REcYcLz-z FLus @As 52 M M coMPREssoR 'a SI/STR'PPER50 /FLuE sAs 40 /HOPPER HOT REcYcLE H2 NET H2 a. u 25\ /49cATALYs'r GAS59 I E; TRANI-sing 7 1n H 25B ."9 54 2o /Jl- J I 53f *L I8 P J PRoDUcT 11 \l \f l I i g lo |7 SYSTEM ElLEJi-:

46 36 27 ,55 /48 :itesm 52 f :i z- 4* T 33 E E; Q--

Si 15:2 55 FUEL R i; E EE 54 -51- 1 1:* :i: 45 44 L 28 1 I 1 tim L :j AI 1i AK BURNER 421 4|B REACTOR K AIR 1 A2 A/ my@ I-Il I\\ /V LuREGENERATED A5 cATALYsT/ l l ,42B

f 2l zg/ .f 2 5 I4 42 o2-FREE T L f- FLUE LIFT @As @j T "N INVENTOR.

ARTHUR L. CONN AT TORNIEY.

United States Patent O FLUID HY DROFORMING WITH CATALYST RECYCLE ArthurL. Conn, Chicago, Ill., assignor to Standard Oii Company, Chicago, Ill.,a corporation of Indiana Application December 1, 1951, Serial No.259,425` 5 Claims. (C1. 196--50) This invention relates to the treatmentof hydrocarbons of the naphtha boiling range with solid catalyst ofsmall particle size in the presence of hydrogen by a iluidized solidstechnique for obtaining octane number improvement, dehydrogenation,aromatization, isomerization, desulfurization, etc., and it pertainsmore particularly to an improved fluid hydroforming system embodyingcatalyst recycle.

Heretofore, fluidized solids technique has been extensively employed foreffecting catalytic cracking of gas oils and to some extent foreifecting reforming of lower boiling hydrocarbons. Hydroforming has beencommercially employed in xed bed units and although it has often beenproposed that uidized solids technique be employed for effectinghydroforming it has `been recognized that hydroforming presents certainproblems that are not encountered in` catalytic cracking. Whereas incatalytic cracking high catalyst to oil ratios may be employed totransfer heat to the reactor and thus supply the endothermic heat ofreaction, in lluid hydroforming they result in a degradation of product`quality and/or yields. The use of higher naphtha preheat temperaturescauses some thermal reforming, with resultant degradation in yield andproduct quality. Increasing the temperature of the recycle gas above1l00 F; presents additional metallurgical problems in furnace design.While cracking catalyst is effective in its fully oxidized state as itleaves the regenerator, hydroforming catalyst requires some degree ofreduction or conditioning. Hydroforming requires high pressures usuallyof the order of 100 to 500 p. s. i. which in turn creates problems ofovercoming erosion and catalyst losses when catalyst-containing fluidsare `depressured. In catalytic cracking, if some hydrocarbon materialfrom the reaction side gets into the regenerator, it merely burnstherein, but if appreciable amounts'of hydrogen should find their wayinto the regenerator, the operation would be more hazardous. An objectof my invention is to provide almethod and means for solving these andotheriproblems presented by the application of luidized solidstechniqueto hydroforming.

Although temperatures are usually remarkably uniform in iuidized solidsbeds, when the diameter of the bed is sufficiently `small as compared toits height and the total heat requirement is supplied at the base of thebed, there is a tendency for-the top of the bed to` be somewhat coolerthan the bottom thereof. `This condition exists in iuid hydroformingbecause it is necessary to have a relatively small diameter `vessel toinsure high enough velocities for iluidization, while a relativelylargeheight is required to give the low space velocities required. It isimpractical to supply additional heat to the upper part of the densephase uidized bed by indirect heat transfer (in the manner that heat isremoved in the regenerator) because materials of construction are notreadily available for withstanding the required operation conditions.`An object of my invention is to provide an improved method and means forobtaining greater uniformity of" temperature in tall, small diameterlluidized beds. A further object is to provide an improved method andmeans for adding heat to such a iluidized solids system.

An important object of my invention is to minimize regenerationrequirements of fluidized catalysts employed for treating hydrocarbonsin the presence of hydrogen. Other objects of the invention will becomeapparent as the detailed description of the invention proceeds.

ln practicing my invention, I employ a tall vertical reactor whoseheight is about 5 to 20 times its diameter. superheated hydrogen,preferably a recycle gas stream, at a temperature of about 1100 to 1300F., is introduced at the base of the reactor and distributed across thecross sectional area thereof. Catalyst may be introduced at the base ofthe reactor with said superheated hydrogen or may be introduced at ahigher level in the reactor either with or without a preconditioning orpartial reduction step. Charging stock vapors are introduced into thereactor with recycled hydrogen or, when the latter is superheated, at alevel spaced from the hydrogen inlet by about .5 to 2 reactor diametersso that below the charging stock inlet the superheated hydrogen willcontact only catalyst and thus will not cause thermal cracking ordegradation of the charging stock. A deep dense turbulent catalyst bedis maintained in the reactor at a depth suicient to give the requiredlow space `velocity which in the case of molybdena-on-alumina catalystsmay be approximately 0.5 w./hr./w., the ratio of hydrogen to chargingstock being at least as great as heretofore employed in fixed bedoperations, e. g. in the range of l:l to 10:1 or about 5:1 on a molbasis.

in order to minimize product degradation, the weight ratio ofregenerated (and reconditioned) catalyst to oil should be in the rangeof .l to 1 or preferably .3` to .4. Catalyst is Withdrawn at acorresponding rate directly from the dense phase at a point spaced fromthe point of catalyst inlet so that short circuiting of regeneratedcatalyst may be avoided. The withdrawn catalyst is then conveyed bysteam or hot hydrogen through an internal or external riser to `anelevated stripper-hopperso that the necessary catalyst head is obtainedfor introducing catalyst to the regenerator. Gas from the top of thestripper-hopper is returned to the dilute phase in the reactor to avoidcatalyst loss and the necessity of extraneous depressuring means.Regenerated catalyst may be introduced directly from the regenerator toa point in the reactor below the charging stock inlet or it may bepicked up with superheated hydrogen for introduction into the base ofthe reactor or it may be preconditioned by treatment with hydrogen at atemperature of approximately 800 to 1100 F. and then returned to thereactor.

One method of obtaining greater uniformity of temperature in the reactoris to withdraw catalyst from upper levels of the dense phase, suspendthe withdrawn catalyst in hot hydrogen and return the suspension to thebottom of the reactor. By this method of operation, cyclic circulationof catalyst through the conversion zone is maintained in addition to theturbulence exhibited by the catalyst in the dense phase itself.Furthermore, the treatment of withdrawn catalyst with hot hydrogenassists in the removal of adsorbed hydrocarbons from said` catalyst andthus prolongs catalyst activity and minimizes regeneration requirements.The cyclic flowof catalyst may also be obtained by circulating catalystwith hot hydrogen from the lower to the upper part of the dense phase sothat there Will be a net downward flow of catalyst in the main portionof the turbulent dense phase.

Instead of or in addition `to the cyclic catalyst flow obtained by useof hot hydrogen, an independent cyclic flow of catalyst may bemaintained by the use of hot high pressure flue gas which issubstantially free from oxygen.

n IaboutgZO-to 100 microns in particular size.

airelnalV Thus catalyst may be. withdrawn, from, an. intermediate levelof the dense catalyst phase in the reactor and conveyed by hotoxygen-free tlue gas to a high level so that in the conveying step theyconveyed:l particlesl are heated by the iiue gas to a temperaturel inthey range of 1000 to l2000 F. Thel heating of the catalyst withoxygen-free flue lgas does not appreciably increase its.v carbon.content nor-cause oxidation of the catalyst so that this method ofheating the catalyst avoids the disadvantage of circulatinglargeramounts of catalyst to the regeneration zone (employing higher catalystto oil ratios). The catalyst thus externally reheated by oxygen-freellue gas can be returncdtoone or more upper parts ol the denselfluidized catalyst bed inthe reactor for supplying additionalendothermic heat of conversion to the reactor and for obtaining moreuniform conversion temperatures therein. The

hot; flue gas separated from catalyst which it has reheated may, belintroduced into the upper part of the regenerator sov that no additionaldepressuringequipment will be required and loss of catalyst from thesystem will be prevented..

Thek inventionrwillbe more clearlyfunderstood from the followingdescription of a specific example thereof read inconjuntion with theaccompanying drawing which isal schematic Flow diagram of a 2000 barrelper day lluid hydroformingunit embodying my invention. The inventionwill be described as applied to a 53 A. P. I. naphthal of low sulphurcontent boiling in the range of about 20G-to 330 F., having a C. F. R.R. octane number ofV about 60 which is hydroformed by the use of amolybdena-on-alumina catalyst having a particle size in the range of lto 200 microns, most of the catalyst being Ther catalyst should be asfree as possible from contaminantssuch as iron, the` preferred catalystbeing made either from pure aluminum chloride by processes of the typedisclosedl in U. S, 2,432,286, 2,481,824, etc., or from high purity pas-Y sive aluminum metal by processes of the type disclosed in U. S.2,274,634-, 2,345,600, 2,371,237, etc. The molybdena content of thecatalyst shouldy preferably be in the range of about to 20weightpercent. No` invention is claimed in the catalyst per se, and it shouldbe understood that any knownhydroforming or hydrofining catalyst maybeemployed. The contact time and operating conditions will be dependent,of course, on the particular catalyst and its activity (as well as onthe charging stock and nature of `results desired);y platinum-containingcatalysts (U. S. 2,479,109-) may require somewhat lower temperatures andlower contact time than molybdena-alumina catalysts. OtherV group VImetal oxides or sulfdes on gamma-alumina supports may be employed eitherwith or without kaddedgroup VIII metal oxides or s ultides such asnickel or cobalt oxides or sulides, examples of mixed catalysts beingthe so-called cobalt-molybdate catalysts, nickeltungstate, etc., withmore or less of the oxygen replaced by sulfur The reactor 10 in thiscase is an insulated pressurevessel with a 5 foot 3 inch I. D. lining inabout a 7 foot I. D. shell about 60 feet tall. Charging stock vapors ata temperature of about 850 toV l000 F., preferably about 900 to 950 F.,are introduced by line 11 through a suitable distributor 12 which isabout 3 to 5 feet above distributor grid 13 and is, generally speaking,about onehalt` to two reactor diameters from the bottom ofV the reactor.A simple method of introducing charging stock vapors is to simplyintroduce them at 3 or more spaced points around the 'circumference ofthe reactor. Superheated hydrogen is introduced at the base of thereactor through line 14 and/ or through line 14owhen saidsuperheatedhydrogen is employed as a carrier gas for returningregenerated catalyst. A turbulent dense phase catalyst bed is maintainedin the reactor with its upper level about 55 feet from the reactorbottom. In this example, the naphtha is supplied at the rate ofapproximately 2000 barrels per day or approximately 22,3750 pounds perhour, 'superheated hydrogen is introduced at the base of the 4 reactorat about 110cm 130.0" E., er. about 1.20.0 E..y at the rate of about1300 mols` per hour, or approximately 18,000 pounds per hour. The upwardgas velocity is thus in the range of about .5 to 1.5, or approximately.8 to .9 foot per second, giving a catalyst bulk density ofapproximately 30 pounds per cubic foot. The charging stock contact timeis approximately one minute and the Weight space velocity is about .4 to.5. The pressure at the base of the reactor is approximately 260 p. s.i. g. and at the top of the reactor about 250p. s. i. g. Reactortemperature is about 800 to 1000 F.

The product stream is withdrawn from the dilute phase in the upper partof the reactor through` cycloneseparator 15, separated catalystparticles being returned by dip leg 16 to a point below the dense phaselevel. From separator 15, the product stream passes by line 17 throughheat exchanger l0 and thence to product separation system 1.9. Since noinvention is claimed in the` product separation system, itrequiresnodetailed description. Generally speaking, however, theA product streamvmaybe further cooled by heat exchange with incoming charging stockv andintroduced into a baled scrubbing section in a fractionating toweroperating4 at substantially reaction` pressure (about 240p. s. i. g.)and with a top temperatureof about 270 and a bottom temperature of about420 l?. Catalyst particles are removed in heavy condensate-and may berecycled to the reactor. The overhead from this fractionator-ma-y thenbe condensed and the recycled hydrogen separated therefrom fromcompression and recycling by line 20 through heat exchanger 18.So-called polymer` may be recovered. from the highV boilingA materialleaving' the first fractionator, the unrecycled'portion ofthe separatedhydrogen may be scrubbed; with incoming` charging stock for recovery ofgasoline boiling range componentsl therein anduncondensed gaseouscomponente andproduct naphtha. maybe separately recovered in any knownmanner. In the system herein described, thevdry gas production may beabout l0 to l2 Weight per cent, polymer about 3 weight per cent,butane-containing gasoline about Weight per cent, and coke approximately.5 weight per cent. The productl gasoline kis substantially free fromsulfur andmay be characterizedbya C.' R. R. octane number ofapproximately 100.

The recycled hydrogen is heated in exchanger 18.l to approximatelyr700F. and then superheated in furnace 21 to a temperature o f about 1100 tol300 Fior return through lines114. or 14a to the base of the reactor.

Catalyst is withdrawn from. the lower part of the reactor above= grid 13through line 22. and carried by a lift gas. suchas steam (or hothydrogen from line 20.) through line 2.4, to an intermediate point inLthe stripper-hopper vesselZS which inv thiscaseis about 2 to.3 feet-indiameter by about 25 feet iniheight, The stripper-hopper may b eprovided with an aeration line inlet. 25a and an aeration distributorgrid 25h so that aeration and/ or stripping may. be effected in vessel25 by the introduction of steam, hot hydrogen or any'other inertA gasthrough` line 25a. Separated gas from the, top of vessel 25 isreturnedby line 26 to the upper part of reactor 10. Construction costs:are minimized by mounting vessel 25 immediately abovel andcontiguouswith the topof reactor 10 so that line26 may be a simpleverticalpipe; however, an external line may be employed for connectingthe top of: the vessel 25- to the top of reactor 10, it beingimportant'that. anyy catalyst in the stripper gases be, recovered andthatdepressuring expense, be minimized.

Catalyst from the-lower partV of stripper-hopper 25 passes by a standpiiy` 27 to a level below the dense catalyst phase inthe regenerator2S. Air-isintroduced into regenerator through line 29 at therate ofaboutA ll14`mols per hour (32,70 pounds per hour) under a pressure ofabout 26,0 p. s. i. g., the pressure above thelevel of grid 3.0 boeingabout 257 p. s. i. g, andlat ther top of the regenerator beingabout'253fp. s. i. g. Itshould; be understood that standpipe 27 is ofsufficient length so that the presv.passed through pressure reducingvalve sure above the control valve at the base thereof is approximately258 p. s; i. g., the pressure at the top of vessel 25 beingsubstantially the same and only slightly above the approximately 250p.`s. i. g. maintained in the top of reactor 10.

To remove the heat `oi" regeneration and thus avoid heat damage tothecatalyst, `cooling tubes 31 may be mounted therein, cooling water beingintroduced from line 32 to separator 33 and thence conveyed by line 34to the base of tubes 31, the top of said tubes being connected by line35 to separator 33 from which steam may be withdrawn through line` 36. Aseries of `filters 37 may be employed in the upper part of theregenerator for avoiding any `catalyst loss and preventing erosiondiftculties when flue gas is reduced to atmospheric pressure. The usualblow-back system may be employed by introducing hot blow-back air orflue gas through line 38 to remove deposits from part of the filterswhile the remaining filters are in operation, the filtered ilue gasbeing 39 and conveyed .by line 40 to a suitable flue.

Since reactor is so tall and narrow, there `is a tendfency for the upperpart of the iiuidized bed to be some- 'what cooler than the lower partthereof. ln order to .maintain a more uniform temperature throughout the.reaction zone, catalyst may be withdrawn from one or vmore upper levelsof the dense'catalyst phase through lline 41 to standpipe 42 and thiswithdrawn catalyst, without regeneration, may then be passed throughbranch Aleg 42a of the standpipe and be picked up by hot hydro- ;gen inline 14 and returned to the base of the reactor. This external recycleof unregenerated catalyst obtains two important advantageous results: itcauses the circulatory flow of catalyst through the reactor which`inturn makes for greater uniformity of temperatures therein and tends tomaintain the activity of the catalyst at a highlevel by the contactingwhich it receives with hot hydrogen in the absence of charging stock.instead of simply returning the withdrawn catalyst with superheatedhydrogen through line 14, the catalyst from the base of branch leg 42amay be picked up from hydrogen from line and then passed through aheating zone thereby increasing the heat capacity ofthe total streamintroduced through line 14. Standpipe 42 may be heated and/or thecatalyst flowing therethrough may be separately heated before beingsuspended in superheated hydrogen from line 14. When an exothermichydrogenation reaction is being effected, the withdrawn and recycledcatalyst stream may be picked up with recycled hydrogen and thereafterbe cooled before its introduction at the base or at other points in thereactor. The external recycle of catalyst from an upper to a lower pointof the reactor by means of recycled hydrogen is thus advantageous frommany standpoints but in the particular example herein described, itschief benefits are the attainment of more uniform conversiontemperatures from top to bottom in the reactor and the prolongedcatalyst activity that is made possible by contact with hydrogen at hightemperature.

Alternatively, more uniform temperatures may be obtained from top tobottom in the reactor by withdrawing catalyst from the dense phasetherein through lines 41a and 41h at intermediate levels in the densephase and employing branch leg 42h instead of 42a of standpipe 42. Ylnthis case a gaseous, liquid or even a solid fuel from line 43 may beburned with air or oxygen-containing lgas from line 44 in high pressureburner 45 under conditions to give a tine gas which is substantiallyfree from oxygen. This gas is mixed with recycle ilue gas from 'line 46to reduce its temperature and thus avoid damage Ato the catalyst, andthe combined gases pass through line 47 to the base of branched leg 42bof standpipe 42 where they pick up catalyst and carry it through line 48to ihot catalyst hopper 49 which is preferably superimposed ,abovereactor 10. Part of the gases from hopper 49 may pass directly to theflue gas compressor 50, or it may be desirable to pass all of the gasesthrough line 51 and filters 37 to remove catalyst fines before return tothe recycle ilue gas compressor 50 by line 52, a small amount ofhydrogen being added to line .52 when necessary to insure its beingoxygen-free. if half the heat of reaction is suppled by this means, itwill be necessary to Supply 2000 to 3000 lbs. per hour of llue gas fromthe burner, to recycle 40,000 to 50,000 lbs. per hour of hue gas throughline 46, and to circulate about 60,000 to 70,000 lbs. per hour ofcatalyst through line 48.

Standpipe 42 and branch leg 42b should be of such length as to provide apressure of about 260 to 262 p. s. i g. immediately above the controlvalve so that a pressure of approximately, but slightly greater than,253 p. s. i. g. may be maintained in the upper part of hot catalysthopper 49. This hopper may be mounted adjacent to stripper-hopper 25.Catalyst and recycle gas from line 48 may be introduced into the densephase of hopper 49 as shown, or through a bottom grid, or the catalystmay be separated from the ilue gas by means of cyclone separators. Whennecessary, a small portion of the recycle` gas may be introduced to thebase of the hopper through line 53, and distributed through grid 54 toinsure proper aeration of the catalyst in the hopper. During its trans-`fer through line 48 and in hopper 49 the catalyst is heated by the hotflue gas to a temperature of the order of 1000 to 1200o F., e. g. aboutl F., and such hot catalyst may be returned in regulated amounts 'byline 55 to the upper part of the dense catalyst phase in the reactor.Part or all of the gases from the top of the hot catalyst hopper may bereturned by line 51 to the upper part of regenerator 23 so that nocatalyst will be lostfrom the system and no separate catalyst recoveryor depressuring equipment will be required. l

By withdrawing catalyst from the intermediate part'of the dense phase inthe long narrow reactor, heating it with oxygen-free due gas in line 4Sand hopper 49 and returning it to the upper part of the reactor by line55 as hereinabove described, additional heat is provided for supplying apart of the heat of reaction and more uni form temperatures areobtainable throughout the dense` phase catalyst bed. lt should beunderstood, of course, that the reheated catalyst may be returned to thereactor through a plurality of lines 5S leading to spaced levels in thedense phase and that similarly catalyst may be withdrawn from variouslevels in the dense phase by lines 41, 41a, 41h, etc., so that thissystem of operation provides great flexibility. lt is altogetherseparate and independent from the regeneration and may be employed forsupplying endothermic heat to reactors employing catalysts which do notrequire regeneration.

When regeneration is employed as in the case of molybdena-on-aluminacatalyst as described in this example, about 8000 pounds per hour ofregnerated catalyst is withdrawn from the lower part of the regeneratorthrough standpipe 56 and picked up by superheated hydrogen from line 14afor return to the reactor. It should be understood, however, that ifdesired the Withdrawn cata` lyst may be partially reduced or conditionedin the separate treating zone before or after it is introduced into thereactor. For example, catalyst from standpipe 56 may be introduced intoa small conditioning vessel and contacted with hot hydrogen from line 20or a mixture of such hot hydrogen with superheated hydrogen from line14. If desired, the off-gases from this preconditioning step may beintroduced above the dense phase level in reactor 10 and the catalystmay be returned directly to the base of the reactor as more fullydescribed in copending application Ser. No. 253,498 led October 27,1951. Many other alternative arrangements and modications will beapparent from the above description to those skilled in the art. It willalso be understood that the specific operating conditions hereinabovedescribed are applicable to the defined charging stock and molybavenue-ii Y depa-ou-alumina catalyst; with other charging stocks and othercatalysts, operating conditions will, of course, `be selected whichtareknown to be effective with such catalysts and charging stocks and sincethese are well known to those skilled in the art, they require nodetailed desciiption. It will be understood, of course, thatstandpipes'such as lines 22, 27, 42u, 42h, 55 and 56 will be providedwith usual slide valves or equipment to control pressure and/or ow fromtheir discharge ends and that suitable aeration lines will be providedabove said valves wherever required for maintaining the catalyst inluidized condition. It will also be understood that a portion of thesuperheated hydrogen `from line 14 may be employed instead of anoxygen-free liuc gas for reheating catalyst and conveying it to anelevated hopper such as hopper 49; when superheated hydrogen is employedfor this purpose, hopper 49 will be maintained at a top pressuresubstantially the same but slightly higher than the top pressure inreactor and the gases from line 51 will be returned to the top of thereactor instead of tothe top of the regenerator. An even simpler methodof distributing the heat to the upper part of the reactor is to providea simple tube riser within the reactor itself and to inject a portion ofthe superheated hydrogen at the base of said riser for carrying catalystto the upper part of the reactor, out of contact with the dense phasetherein while it is being heated with the supcrheated hydrogen stream;however, this method would require heating the naphtha and/ or thehydrogen to much higher temperatures to supply the heat of reaction. lnall cases it will be noted that the temperature in the reactor is mademore uniform by withdrawing catalyst directly from the dense phase,heating the withdrawn catalyst by direct contact with an oxygen-freeheating gas and returning the heated catalyst to the dense phase in theconversion zone to elect cyclic flow of catalyst through the conversionzone in addition to the turbulent ilow therein.

l claim:

1. ln an en dothermic process for the conversion of a hydrocarboncharging stock of the naphtha boiling range *in the presence of hydrogenby means of iluidized solids conversion catalyst in a high pressureZoneV which is so narrow in cross sectional area with respect to itsheight that the upper part of said zone tends to be cooler than thelower part because of the endothermic nature of the process, theimproved method of operation which cornprises supplying heat to thelower part of. said Zone by introducing thereto a stream of hydrogenwhich has been superheated to a temperature above 1000" F., withdrawingdense phase fluidized catalyst directly from an intermediate level inthe reaction zone, suspending said withdrawn catalyst in ahotoxygen-free gas of the class consisting of flue gas, hydrogen andmixtures thereof at a higher temperature than the temperature of thewithdrawn catalyst, conveying said withdrawn catalyst to a higherelevation than the reaction zone by means of said hot oxygen-free gas,separating said hot oxygen-free gas from catalyst which has been heatedthereby and returning the heatedcatalyst, to the upper part of saidzone.

2. The methodV of effecting fluid hydroforrning in a column of lluidizedhydroforming catalyst whose depth is Iat leasttve times its diameter,which method comprises introducing 'hot regenerated catalyst vand hothydrogen at the base of said column for supplying heat at Vthe lowerpart of the column, withdrawing catalyst from an intermediate level insaid column, suspending said withdrawn catalyst in oxygen-free liuc gaswhich is at a higher temperature than that of the Withdrawn catalyst andconveying said withdrawn catalyst in said hot oxygenree iiue gas to a'higher level than that of the conversion zone, whereby the catalyst isheated in transit, separating from heated catalyst and returning theheated catalyst at the upper part of said column for supplying heattosaid upper part'.

3. The method of effecting endothermic fluid hydro*- Eorming in afluidized catalyst bed of which the yheight is at least 5 times thediameter, which method comprises introducing at the base of said bed aVhydrogen stream which has been superheated to a temperaturesubstantially higher than 1000e F., introducing at a level spaced fromthe base of the bed vapors of a naphtha charging stock which have beenpreheated to at least a conversion temperature in the range of 850 to950 F., passing the hydrogen and naphtha vapors upwardly through saidbed at a velocity to maintain it in tluidized condition at saidconversion temperature and under a pressure in the range of tov 500 p.s. i. g. whereby the top of the bed tends to become cooler than thebottom thereof, continuously removing a dense phase stream of catalystfrom said'bed to a stripping zone and then to a regeneration zone,regenerating the catalyst with an oxygen-containing gas and returningregenerated catalyst to said bed, continuously removing a second densephase stream of catalyst from the bed at a level spaced from thecharging stock inlet, suspending said second catalyst stream in'anon-oxidizing gas of the class consisting of hydrogen, flue gas andmixtures thereof which non-oxidizing gas has been preheated to atemperature substantially above 10'00" F., conveying said suspendedcatalyst upwardly in said nonoxidizing gas while effecting heat transferfrom said gas to said catalyst to raise the temperature of the catalystto at least about 1000" F. and returning said heated catalyst at'theupper part of the luidized catalyst bed.

4. The method of claim v3 wherein the second catalyst stream iswithdrawn from an intermediate level in the catalyst bed, conveyed to ahigh level by'hot flue gas and returned to the upper part of thecatalyst bed after separation of flue gas therefrom.

5. The method of claim 4 which includes the step of recycling a portionof the separated ue gas to hot flue gas employed for suspending,elevating and heating the catalyst.

Refeii't-enees,v Cited in the file of this patent UNlTED STA-TES PATENTS2,409,690 Nicholson et al. Oct. 22, 1946 2,459,836 Murphree Jan. 25,1949 2,471,064 Hall et al May 24, 1949 2,472,844 Munday et al June 14,1949 2,643,214 Hartwig lune 23, 1953

1. IN AN ENDOTHERMIC PROCESS FOR THE CONVERSION OF A HYDROCARBONCHARGING STOCK OF THE NAPHTHA BOILING RANGE IN THE PRESENCE OF HYDROGENBY MEANS OF FLUIDIZED SOLIDS CONVERSION CATALYST IN A HIGH PRESSURE ZONEWHICH IS SO NARROW IN CROSS SECTIONAL AREA WITH RESPECT TO ITS HEIGHTTHAT THE UPPER PART OF SAID ZONE TENDS TO BE COOLER THAN THE LOWER PARTBECAUSE OF THE ENDOTHERMIC NATURE OF THE PROCESS, THE IMPROVED METHOD OFOPERATION WHICH COMPRISES SUPPLYING HEAT TO THE LOWER PART OF SAID ZONEBY INTRODUCING THERETO A STREAM OF HYDROGEN WHICH HAS BEEN SUPERHEATEDTO A TEMPERTURE ABOVE 1000* F., WITHDRAWING DENSE PHASE FLUIDIZEDCATALYST DIRECTLY FROM AN INTERMEDIATE LEVEL IN THE REACTION ZONE,SUSPENDING SAID WITHDRAWN CATALYST IN A HOT OXYGEN-FREE GAS OF THE CLASSCONSISTING OF FLUE GAS, HYDROGEN AND MIXTURES THEREOF AT A HIGHERTEMPERATURE THAN THE TEMPERATURE OF THE WITHDRAWN CATALYST, CONVEYINGSAID WITHDRAWN CATALYST TO A HIGHER ELEVATION THAN THE REACTION ZONE BYMEANS OF SAID HOT OXYGEN-FREE GAS, SEPARATING SAID HOT OXYGEN-FREE GASFROM CATALYST WHICH HAS BEEN HEATED THEREBY AND RETURNING THE HEATEDCATALYST, TO THE UPPER PART OF SAID ZONE.